Butadiene extraction process

ABSTRACT

A process for recovering 1,3-butadiene from a C 4  fraction, where the butadiene extraction processes may be operated at an intermediate pressure using a liquid ring type compressor. The use of a liquid ring compressor, among other process options presented herein, may advantageously reduce capital and operating costs, similar to the compressorless option, while mitigating the risks associated with the higher operating temperatures and pressures associated with the compressorless option. Thus, the embodiments of the processes disclosed herein encompass the best features of the conventional design (low pressure, with a compressor) with the advantages of the compressorless design (low capital and operating cost), as well as other advantages unique to the systems disclosed herein.

CROSS-REFERENCE TO RELATED APPLICATIONS

This application, pursuant to 35 U.S.C. §119(e), claims priority to U.S.Provisional Application Ser. No. 61/720,038, filed Oct. 30, 2012, whichis herein incorporated by reference in its entirety.

FIELD OF THE DISCLOSURE

Embodiments disclosed here relate to recovery of butadiene from a mixedhydrocarbon stream. More specifically, embodiments disclosed hereinrelate to an improved butadiene extraction process wherein the degasseroperates at an intermediate pressure.

BACKGROUND

Butadiene is an important base chemical and is used, for example, toprepare synthetic rubbers (butadiene homopolymers,styrene-butadiene-rubber or nitrile rubber) or for preparingthermoplastic terpolymers (acrylonitrile-butadiene-styrene copolymers).Butadiene is also converted to sulfolane, chloroprene and1,4-hexamethylenediamine (via 1,4-dichlorobutene and adiponitrile).Dimerization of butadiene also allows vinylcyclohexene to be generated,which can be dehydrogenated to form styrene.

Butadiene can be prepared from saturated hydrocarbons by refiningprocess or by thermal cracking (steam cracking) processes, in which casenaphtha is typically used as the raw material. In the course of refiningor steam cracking of naphtha, a mixture of methane, ethane, ethene,acetylene, propane, propene, propyne, allene, butenes, butadiene,butynes, methylallene, C₄ and higher hydrocarbons are obtained.

Owing to the small differences in the relative volatilities of thecomponents of a C₄ cut, obtaining 1,3-butadiene from the C₄ cut is acomplicated distillation problem. Therefore, the separation is carriedout by extractive distillation, i.e. a distillation with addition of anextractant which has a higher boiling point than the mixture to beseparated and which increases the differences in the relativevolatilities of the components to be separated. The use of suitableextractants allows a crude 1,3-butadiene fraction to be obtained fromthe C₄ cut mentioned by means of extractive distillation, and saidfraction is subsequently further purified in purifying distillationcolumns.

The butadiene recovery processes typically use 3- or 4-column extractivedistillation systems to separate a mixed C₄ stream into productfractions, including a lights/butane/butenes stream (Raffinate-1product), a crude butadiene product, which may be sent to a conventionaldistillation system for further purification, and C₃ acetylenes(propyne) and C₄ acetylenes streams, which may be sent to a selectivehydrogenation unit, for example.

In the present context, crude 1,3-butadiene refers to a hydrocarbonmixture which has been obtained from a C₄ cut from which at least 90% byweight of the sum of butanes and butenes, preferably at least 98% byweight of the sum of butanes and butenes, more preferably at least 99%by weight of the sum of butanes and butenes, and simultaneously at least90% by weight of the C₄ acetylenes, preferably at least 96% by weight ofthe C₄ acetylenes, more preferably at least 99% by weight of the C₄acetylenes, has been removed. Crude 1,3-butadiene contains the1,3-butadiene product of value frequently in a proportion of at least80% by weight, preferably 90% by weight, more preferably more than 95%by weight, remainder impurities. Accordingly, pure 1,3-butadiene refersto a hydrocarbon mixture which contains the 1,3-butadiene product ofvalue in a proportion of at least 98% by weight, preferably of at least99.5% by weight, more preferably in the range between 99.7 and 99.9% byweight, remainder impurities.

Typical processes to recover butadiene from mixed C₄ streams includeextractive distillation processes, which may incorporate use ofselective solvents. Examples of extractive distillation processes arefound, for example, in U.S. Pat. Nos. 7,692,053, 7,393,992, 7,482,500,7,226,527, 4,310,388, and 7,132,038, among others.

The extractive distillation processes described in the above mentionedpatents typically fall into one of two categories, a conventional lowpressure process including a compressor or a high pressure“compressorless” process, such as disclosed in U.S. Pat. No. 7,692,053.

The compressorless design has the advantages of lower capital costs, asthis design option eliminates the recycle gas compressor entirely.However, there are several disadvantages. For example, for thecompressorless design, the degasser may be operated at an overheadpressure of about 4.21 kg/cm² gage, slightly above the extractivedistillation system (including the main washer, rectifier andafterwasher) pressure. Consequently, the degasser operates atcorrespondingly higher temperatures: about 148° C. at the top of thedegasser and about 193° C. at the bottom of the degasser. In contrast,the degasser in the conventional design may be operated at an overheadpressure of only 0.7 kg/cm² gage, and at much lower temperatures: about105° C. at the top of the degasser and about 149° C. at the bottom ofthe degasser.

The roughly 44° C. hotter degasser temperatures for the compressorlessdesign results in two distinct disadvantages. First, vinyl cyclohexene(VCH, or butadiene dimer) make increases with increasing temperature anda higher dimer make results in lower yield and potentially higherequipment fouling rates. Second, there is a potential for greater riskdue to having high C₄ acetylene concentrations at the higher operatingtemperatures and pressures. To mitigate this risk, the vinyl acetyleneconcentration in the degasser must be kept lower (below 20 mol. %).However, limiting the vinyl acetylene concentration may lead toadditional 1,3-butadiene losses, and thus lower yield.

SUMMARY OF THE CLAIMED EMBODIMENTS

It has now been found that butadiene extraction processes may beoperated at an intermediate pressure using a liquid ring typecompressor. The use of a liquid ring compressor, among other processoptions presented herein, may advantageously reduce capital andoperating costs, similar to the compressorless option, while mitigatingthe risks associated with the higher operating temperatures andpressures associated with the compressorless option. Thus, theembodiments of the processes disclosed herein encompass the bestfeatures of the conventional design (low pressure, with a compressor)with the advantages of the compressorless design (low capital andoperating cost), as well as other advantages unique to the systemsdisclosed herein.

In one aspect, embodiments disclosed herein relate to a process forrecovering 1,3-butadiene from a C₄ fraction. The process may include:feeding a hydrocarbon fraction containing butanes, butenes,1,2-butadiene, 1,3-butadiene, C₄ acetylenes, C₃ acetylenes, and C₅₊hydrocarbons to an extractive distillation system; contacting thehydrocarbon fraction with a solvent in the extractive distillationsystem to selectively dissolve a portion of the hydrocarbon fraction;recovering a vapor fraction comprising a first portion of the butanesand the butenes from the extractive distillation system; recovering anenriched solvent fraction comprising the 1,3-butadiene, the1,2-butadiene, C₄ acetylenes, C₃ acetylenes, C₅₊ hydrocarbons, and asecond portion of the butanes and the butenes; feeding the enrichedsolvent fraction to a rectifier to at least partially degas the enrichedsolvent; recovering a second portion of the butanes and butenes from therectifier as an overheads fraction; recovering the C₃ and C₄ acetylenes,1,3-butadiene, 1,2-butadiene, and C₅₊ hydrocarbons from the rectifier asa side draw fraction; recovering a partially degassed solvent comprising1,2-butadiene and C₄ acetylenes from the rectifier as a bottomsfraction; feeding at least a portion of the partially degassed solventto a degasser to further degas the solvent; recovering an overheadsfraction comprising at least one of C₄ acetylenes and 1,2-butadiene fromthe degasser; recovering a side draw fraction comprising the C₄acetylenes from the degasser; recovering a bottoms fraction comprisingdegassed solvent from the degasser; compressing the degasser overheadsfraction using a liquid ring compressor; and recycling at least aportion of the compressed degasser overheads fraction to the rectifier.

In another aspect, embodiments disclosed herein relate to a system forrecovering 1,3-butadiene from a C₄ fraction. The system may include: aflow conduit for feeding a hydrocarbon fraction containing butanes,butenes, 1,2-butadiene, 1,3-butadiene, C₄ acetylenes, C₃ acetylenes, andC₅₊ hydrocarbons to an extractive distillation system; the extractivedistillation system for contacting the hydrocarbon fraction with asolvent in the extractive distillation system to selectively dissolve aportion of the hydrocarbon fraction; a flow conduit for recovering avapor fraction comprising a first portion of the butanes and the butenesfrom the extractive distillation system; a flow conduit for recoveringan enriched solvent fraction comprising the 1,3-butadiene, the1,2-butadiene, C₄ acetylenes, C₃ acetylenes, C₅₊ hydrocarbons, and asecond portion of the butanes and the butenes; a flow conduit forfeeding the enriched solvent fraction to a rectifier; the rectifier forat least partially degassing the enriched solvent; a flow conduit forrecovering a second portion of the butanes and butenes from therectifier as an overheads fraction; a flow conduit for recovering the C₃and C₄ acetylenes, 1,3-butadiene, 1,2-butadiene, and C₅+ hydrocarbonsfrom the rectifier as a side draw fraction; a flow conduit forrecovering a partially degasses solvent comprising 1,2-butadiene and C₄acetylenes from the rectifier as a bottoms fraction; a flow conduit forfeeding at least a portion of the partially degassed solvent to adegasser; the degasser for further degassing the solvent; a flow conduitfor recovering an overheads fraction comprising at least one of C₄acetylenes and 1,2-butadiene from the degasser; a flow conduit forrecovering a side draw fraction comprising the C₄ acetylenes from thedegasser; a flow conduit for recovering a bottoms fraction comprisingdegassed solvent from the degasser; a liquid ring compressor forcompressing the degasser overheads fraction; and a flow conduit forrecycling at least a portion of the compressed degasser overheadsfraction to the rectifier.

Other aspects and advantages will be apparent from the followingdescription and the appended claims.

BRIEF DESCRIPTION OF DRAWINGS

FIG. 1 is a simplified flow diagram of a process for butadiene recoveryaccording to embodiments disclosed herein.

FIG. 2 is a simplified flow diagram of a process for butadiene recoveryaccording to embodiments disclosed herein.

FIG. 3 is a simplified flow diagram of a process for butadiene recoveryaccording to embodiments disclosed herein.

FIG. 4 is a simplified flow diagram of a process for butadiene recoveryaccording to embodiments disclosed herein.

As noted, the flow diagrams in FIGS. 1-4 are simplified, and do notillustrate pumps, valves, control valves, filters, reboilers,condensers, and other equipment commonly associated with distillationcolumns and general petrochemical operations, and these would beunderstood to be present by one skilled in the art based on the Figuresand the Detailed Description provided below.

FIG. 5 is a comparison of vaporization percentages versus reboiler dutyas a function of temperature for two different pressures.

DETAILED DESCRIPTION

Embodiments disclosed here relate to recovering butadiene from mixed C₄hydrocarbon streams. More specifically, embodiments disclosed hereinrelate to improving the operations and economics of butadiene extractionprocesses via use of intermediate pressures and a liquid ring typecompressor.

The C₄ fraction to be used as starting mixture in the present processesis a mixture of hydrocarbons having predominantly four carbon atoms permolecule. C₄ fractions are obtained, for example, in the preparation ofethylene and/or propylene by thermal or catalytic cracking of apetroleum fraction, such as liquefied petroleum gas, light naphtha orgas oil. C₄ fractions may also be obtained by the catalyticdehydrogenation (oxidative and/or non-oxidative dehydrogenation) ofn-butane and/or n-butene. The resulting C₄ fractions generally includebutanes, n-butene, isobutene, 1,3-butadiene and small amounts of C₃ andC₅ hydrocarbons, including methylacetylene, as well as butynes, inparticular 1-butyne (ethylacetylene) and butenyne (vinylacetylene). The1,3-butadiene content is generally from 5 to 80% by weight. For example,a cracker or a CATADIENE unit may contain 15 to 17% butadiene, byweight. Other mixed C₄ feed streams may contain greater or lesseramounts of butadiene. When present in the mixed feed stream,vinylacetylene may be selectively hydrogenated to the desired1,3-butadiene product prior to feed of the mixed C₄ stream to thebutadiene extraction unit. In some embodiments, the mixed C₄ hydrocarbonstream may be provided, for example, by at least one of cracking,oxidatively dehydrogenating, and non-oxidatively dehydrogenating a C₄hydrocarbon stream comprising butane in one or more dehydrogenationreactors to produce a product gas stream comprising butane, butene, andbutadiene.

The above-described hydrocarbon fraction, containing butanes, butenes,1,2-butadiene, 1,3-butadiene, C₄ acetylenes, C₃ acetylenes, and C₅₊hydrocarbons, is fed to a butadiene extraction unit for separation andrecovery of the various hydrocarbons, including one or morelights/butanes/butenes fractions (commonly referred to as a Raffinate-1product), a 1,3-butadiene fraction, a C₃ acetylenes (propyne) fraction,a C₄ acetylenes fraction, which may include a portion of the1,2-butadiene, and a heavies fraction, which may include a portion ofthe 1,2-butadiene and the C₅₊ hydrocarbons. In some embodiments, dimersof butadiene may be formed upstream of the butadiene extraction unit orduring processing of the hydrocarbon fraction within the butadieneextraction unit. The vinylcyclohexene components may be recovered withthe heavies fraction, or may be recovered as a separate fractioncontaining vinylcyclohexene.

It has been found that butadiene extraction processes may be improvedvia use of a liquid ring type compressor for compressing at least aportion of the overheads from the degasser. Referring now to FIG. 1, asimplified process flow diagram for recovering butadiene according toembodiments disclosed herein is illustrated. A mixed hydrocarbon feed 2,including hydrocarbons such as butanes, butenes, 1,2-butadiene,1,3-butadiene, methyl acetylene, vinyl acetylene, and C₅+ hydrocarbons,may be fed to a feed vaporization system (not shown) to vaporize themixed hydrocarbon feed. The vaporized feed is then fed to main washcolumn 44. In main wash column 44, the vaporized feed is contacted witha solvent, and the butanes and butenes are separated from the moresoluble 1,3-butadiene, 1,2-butadiene, methyl acetylene, vinyl acetylene,and C₅₊ hydrocarbons.

Solvents useful in the process as illustrated in FIG. 1 may includebutyrolactone, nitriles such as acetonitrile, propionitrile,methoxypropionitrile, ketones such as acetone, furfural,N-alkyl-substituted lower aliphatic amides such as dimethylformamide,diethylformamide, dimethylacetamide, diethylacetamide,N-formylmorpholine, N-alkyl-substituted cyclic amides (lactams) such asN-alkylpyrrolidones, especially N-methylpyrrolidone (NMP). In someembodiments, alkyl-substituted lower aliphatic amides orN-alkyl-substituted cyclic amides, dimethylformamide, acetonitrile,furfural or NMP are used.

In some embodiments, it is also possible to use mixtures of theseextractants with one another, for example of NMP and acetonitrile,mixtures of these extractants with cosolvents and/or tert-butyl ethers,e.g. methyl tert-butyl ether, ethyl tert-butyl ether, propyl tert-butylether, n- or isobutyl tert-butyl ether. In other embodiments, NMP may bein aqueous solution, with from 0 to about 20 weight % water, or withfrom 7 to 10 weight % water, or with 8 to 8.5 weight % water in otherembodiments.

The butanes and butenes are recovered from main wash column 44 as anoverheads fraction 8 (Raffinate 1). The enriched solvent, including thedissolved hydrocarbons, is recovered from main wash column 44 as abottoms fraction 46.

Bottoms fraction 46 is then fed to rectifier 48 to at least partiallydegas the enriched solvent. Any dissolved butanes and butenes, as wellas other light components may be recovered from rectifier 48 as anoverheads fraction 50, which may recycled for re-processing in main washcolumn 44. Methyl acetylene and butadienes, including both 1,2-butadieneand 1,3-butadiene, and C₅+ hydrocarbons may be recovered from rectifier48 as a side draw 52, and a degassed solvent, which may contain variousC₄ components including 1,2-butadiene, 1-butyne, and vinyl acetylene,may be recovered from rectifier 48 as a bottoms fraction 54.

Bottoms fraction 54 may be fed to a degasser 56, for separation of thesolvent, entrained C₄ components, and a C₄ acetylene fraction, which mayalso include 1,2-butadiene. The C₄ vapors may be recovered from degasserand cooling column 56 as an overheads fraction 58, which may becompressed via liquid ring compressor 60.

Liquid ring compressor 60 serves two functions: compression of thedegasser overhead fraction and cooling of the compressed gas beforerecycle to the rectifier 48. Following compression, a portion of thecompressed gases may be recycled to rectifier 48. In some embodiments,the compressed degasser overhead fraction may be recovered via flow line88 and fed to a separator 90 to separate any condensed gases. The vaporfraction recovered from separator 90 may then be recycled via flow line92 to rectifier 48. The condensate fraction may be recovered fromseparator 90 via flow line 94, at least a portion of which may be cooledvia heat exchanger 96 and recycled to liquid ring compressor 60.

A vinyl acetylene fraction may be withdrawn from degasser 56 as a sidedraw fraction 62, washed with water fed via line 64 in acetylene washer66, and recovered as vinyl acetylene fraction 12. The degassed solventmay be recovered from degasser 56 as a bottoms fraction 68 for recycleand feed to main wash column 44 and afterwash column 70, where thehydrocarbons in the side draw fraction 52 may be separated from thesolvent. Solvent may be recovered from afterwash column 70 as a bottomsfraction 72 and recycled to rectifier 48, and a crude butadiene productstream may be recovered from afterwash column 70 as an overheadsfraction 74.

The crude butadiene product (overheads fraction 74) leaves theextractive distillation section and is then fed to a methyl acetylenedistillation column 76, where methyl acetylene is recovered as anoverheads fraction 10. The bottoms fraction 78 contains the1,3-butadiene, 1,2-butadiene, and heavier hydrocarbons, and is fed tobutadiene fractionator 80. 1,3-Butadiene having a purity of greater than99.6% is recovered from butadiene column 80 as an overheads fraction 6,and the 1,2-butadiene and heavies are recovered as a bottoms fraction14.

In some embodiments, it may be desired to hydrogenate acetylenes infractions 10, 12 to produce additional olefins and dienes. Additionallyor alternatively, it may be desired to use a green oil column to recoveroligomers of butadiene (vinyl cyclohexane) and oligomers of otherolefinic components in the hydrocarbon feed stream that may be producedduring the separations noted above.

Referring now to FIG. 2, a simplified flow diagram for a process forrecovering 1,3-butadiene according to embodiments disclosed herein isillustrated, where like numerals represent like parts. In thisembodiment, the bottoms fraction recovered from the rectifier 48 viaflow stream 54 may be heated via indirect heat exchange prior to feed tothe degasser 56, such as via heat exchanger 98. The heating of therectifier bottoms may vaporize a portion of the remaining dissolvedgasses, such as 1,2-butadiene or C₄ acetylenes. Prior to feeding of theheated rectifier bottoms to the degasser, the heated bottoms may be fedto a degasser feed drum 16 to phase separate the vaporized portion fromthe liquid portion of the effluent recovered from heater 98. A liquidphase may then be recovered from feed drum 16 and fed via flow line 18to degasser 56 and processed as described above with respect to FIG. 1.The vapor phase recovered from feed drum 16 via flow line 40 may then becombined with the compressed vapor fraction of flow line 92 for recycleto rectifier 48.

Degasser feed drum 16 may operate at a pressure slightly above therectifier 48 bottoms pressure, allowing the vapor phase recovered fromfeed drum 16 to flow freely back to rectifier 48 without the need forvapor recompression. Some of the heat input added via exchanger 98 isthus returned to rectifier 48 in the form of the flashed vapors. Therecycle of gas from drum 16 to rectifier 48 at a slightly elevatedtemperature may thus add heat to rectifier 48, and may result inadditional pre-degassing in the bottom section of rectifier 48,contributing to an overall lower degassing requirement in degasser 56.

Degasser feed drum 16 may be a separate vessel, or as illustrated, maybe integral with the degasser 56, forming a single vessel structure.Integrating the feed drum and the degasser into a single vessel mayreduce capital costs. By locating feed drum 16 above or on top ofdegasser 56, the liquid phase in the feed drum may easily flow into thetop of the degasser without the need for additional pumps. Part of theheat input from exchanger 98 thus also flows to degasser 56 in the formof sensible heat contained in the un-flashed liquid fed to degasser 56via flow line 18.

Overall, the use of exchanger 98 and phase separation in drum 16 mayprovide for two stages of flashing, in feed drum 16 and degasser 56,where feed drum 16 may be operated at a pressure greater than that ofdegasser 56. Use of the two stage separations may result in moreefficient C4 degassing, improving separations of the C4 hydrocarbonsfrom the solvent. Further, dissolved gases degassed in feed drum 16 andrecovered via flow line 40 are at a higher pressure level, and do notrequire recompression for recycle to rectifier 48.

Degasser 56 may be operated at a pressure lower than rectifier 56, butonly the gases recovered via flow line 58 require recompression. As aresult, liquid ring compressor 60 may be sized to account for thereduced vapor flow and reduced compression requirements, resulting inlower capital and operating expenses. For example, in some embodiments,a ratio of the vapor flow rate of the compressed degasser overheads(stream 92) to the vapor flow rate of the vapor phase recovered fromdegasser feed drum 16 (stream 40) may be in the range from about 0.1:1to about 1:1; in the range from about 0.2:1 to about 0.8:1 in otherembodiments; and from about 0.25:1 to about 0.5:1 in yet otherembodiments.

Referring now to FIG. 3, a simplified flow diagram for a process forrecovering 1,3-butadiene according to embodiments disclosed herein isillustrated, where like numerals represent like parts. In thisembodiment, the bottoms fraction recovered from the rectifier 48 viaflow stream 54 may be heated via indirect heat exchange prior to feed tothe degasser 56, such as via heat exchangers 4, 98, to vaporize aportion of the remaining dissolved gasses, such as 1,2-butadiene or C₄acetylenes. Heat exchanger 4 may be used to recover heat from thebottoms fraction 68 recovered from degasser 56. Heat exchanger 98 maythen be used further heat the rectifier bottoms before processing of thepartially vaporized rectifier bottoms in degasser feed drum 16 andprocessed as described above with respect to FIG. 2.

Referring now to FIG. 4, a simplified flow diagram for a process forrecovering 1,3-butadiene according to embodiments disclosed herein isillustrated, where like numerals represent like parts. In thisembodiment, the liquid portion recovered from degasser feed drum 16 viaflow line 18 is heated via indirect heat exchange in heat exchanger 30to provide addition heat to the degasser feed, which is then processedas described above.

As discussed above with respect to FIGS. 2-4, the rectifier bottoms maybe heated via indirect heat exchange using exchangers 4, 98. In someembodiments, exchanger 98 may use a heat exchange medium, such as water,steam, or a synthetic organic heat transfer fluid, such as DOWTHERM orothers as may be known to those in the art. It may also be desirable tosuppress vaporization in the heat exchangers and associated piping,favoring vaporization in feed drum 16 or degasser 56 to minimize orprevent fouling. Thus, in some embodiments, heat exchangers 4, 98 may besuppressed vaporization heaters.

Degasser feed drum 16 may be operated at a pressure in the range fromabout 3.5 kg/cm² gage to about 5.5 kg/cm² gage in some embodiments; inthe range from about 4 kg/cm² gage to about 5 kg/cm² gage in otherembodiments; and from about 4.25 kg/cm² gage to about 4.75 kg/cm² gagein yet other embodiments, such as about 4.5 kg/cm² gage. Degasser feeddrum 16 may be operated at a temperature in the range from about 110° C.to about 150° C. in some embodiments; in the range from about 120° C. toabout 140° C. in other embodiments; and in the range from about 125° C.to about 135° C. in yet other embodiments, such as about 130° C.

Degasser 56 may be operated at a pressure in the range from about 1.5kg/cm² gage to about 3.5 kg/cm² gage in some embodiments; in the rangefrom about 2 kg/cm² gage to about 3 kg/cm² gage in other embodiments;and from about 2.25 kg/cm² gage to about 2.75 kg/cm² gage in yet otherembodiments, such as in the range from about 2.3 kg/cm² gage to about2.5 kg/cm² gage. Degasser 56 may be operated at an overhead temperaturein the range from about 100° C. to about 150° C. in some embodiments; inthe range from about 110° C. to about 140° C. in other embodiments; andin the range from about 120° C. to about 130° C. in yet otherembodiments, such as about 125° C. Degasser 56 may be operated at abottoms temperature in the range from about 150° C. to about 200° C. insome embodiments; in the range from about 160° C. to about 190° C. inother embodiments; and in the range from about 170° C. to about 180° C.in yet other embodiments, such as about 175° C.

Rectifier 48 may be operated at a pressure in the range from about 3kg/cm² gage to about 5 kg/cm² gage in some embodiments; in the rangefrom about 3.5 kg/cm² gage to about 4.5 kg/cm² gage in otherembodiments; and from about 4 kg/cm² gage to about 4.5 kg/cm² gage inyet other embodiments, such as in the range from about 4.1 kg/cm² gageto about 4.2 kg/cm² gage. Rectifier 48 may be operated at an overheadtemperature in the range from about 40° C. to about 90° C. in someembodiments; in the range from about 50° C. to about 80° C. in otherembodiments; and in the range from about 60° C. to about 70° C. in yetother embodiments, such as in the range from about 63 to about 68° C.Rectifier 48 may be operated at a bottoms temperature in the range fromabout 60° C. to about 120° C. in some embodiments; in the range fromabout 70° C. to about 110° C. in other embodiments; and in the rangefrom about 75° C. to about 100° C. in yet other embodiments, such as inthe range from about 80° C. to about 95° C.

Heat may be supplied to the rectifier via indirect heat exchange in areboiler using a heating medium having a temperature of less than 130°C. For example, the heating medium used to heat the rectifier reboilermay have an operating temperature in the range from about 80° C. toabout 130° C. in some embodiments; in the range from about 90° C. toabout 125° C. in other embodiments; and in the range from about 100° C.to about 120° C. in yet other embodiments. In some embodiments, the heatexchange medium used in the rectifier reboiler may be controlled suchthat the process-side temperature increase across the reboiler is in therange from about 5° C. to about 15° C.; and in the range from about 8°C. to about 12° C. in other embodiments, such as a delta of about 10° C.

The two-stage degassing provided for in the degasser feed drum 16 andthe degasser 56, as well as heat introduced to rectifier 56 via vaporstreams 40, 92 may allow the rectifier reboiler to operate at a lowpercent vaporization. For example, in some embodiments, the rectifierreboiler may operate having a percent vaporization across the reboilerin the range from about 3 wt. % to about 9 wt. %; in the range fromabout 4 wt. % to about 8 wt. % in other embodiments; and in the rangefrom about 5 wt. % to about 7 wt. % in yet other embodiments, such as inthe range from about 6 wt. % to about 6.5 wt %. The combination of lowpercent vaporization and low temperatures (both hot side and cold side)may significantly reduce fouling in the rectifier reboiler.Additionally, the low percent vaporization and reduced fouling maypermit the rectifier reboiler to be a conventional type heat exchanger,including single-pass heat exchangers, as opposed to a suppressedvaporization type exchanger.

EXAMPLE

A process for recovering butadiene according to embodiments disclosedherein, similar to that illustrated in FIG. 3, is compared to aconventional process for recovering butadiene (using a screw type orcentrifugal compressor as well as a cooling column) and a compressorlessprocess (also including a cooling column) for recovering butadiene,using the following conditions.

For the compressorless design, the degasser is operated at an overheadpressure of 4.21 kg/cm² gage, slightly above the extractive distillationsystem (main washer, rectifier and afterwasher) pressure. Consequently,the degasser operates at correspondingly higher temperatures: 148° C. attop and 193° C. at bottom.

For the conventional process, the degasser operates at an overheadpressure of only 0.7 kg/cm² gage, and at much lower temperatures: 105°C. at top and 149° C. at bottom.

For this example, the embodiment as illustrated in FIG. 3 uses aonce-through, co-current rectifier reboiler that uses partially cooleddegasser bottoms (lean solvent) on the shell side. Partial vaporization(degassing) occurs in the reboiler tubes, and the vapor/liquid mixtureis heated to 90° C. The partially degassed rich solvent at 90° C. isthen pumped by a degasser feed pump to subsequent exchangers where it isfurther heated in suppressed vaporization type exchangers. The degasserfeed pump provides sufficient discharge pressure to ensure that novaporization (degassing) occurs in any of the exchangers. The firstexchanger is the tube side of the degasser feed/effluent exchanger,where the rich solvent is heated up to approximately 138° C. on the tubeside. The degasser feed/effluent exchanger is a two shell exchangerbecause of the large temperature cross that occurs in this exchanger.Degasser bottoms (lean solvent) at 175° C. are used as the heatingmedium on the shell side of the exchanger. The lean solvent is cooled to120° C. in the degasser feed/effluent exchanger before being sent to theshell side of the rectifier reboiler and subsequently to the butadienecolumn reboiler, feed vaporizer, and solvent cooler.

Heated rich solvent from the degasser feed/effluent exchanger is thensent to the degasser feed heater where the rich solvent is furtherheated against low pressure steam (150° C.) to a temperature ofapproximately 138° C. Rich solvent at its final temperature is thenflashed across a control valve into the degasser feed drum, which sitson top of the degasser. The feed drum rides off of the rectifier bottomspressure, and the flashed gas flows freely back to the rectifier whereit enters below the bottom bed. Un-flashed liquid from the degasser feeddrum then flows by pressure difference/gravity into the top of thedegasser where additional feed flashing occurs. The partially degassedsolvent then flows down the multi-bed degasser, where essentially all ofthe remaining C₄ hydrocarbons are completely stripped from the solvent.Stripping heat is provided by the Degasser reboiler, utilizing mediumpressure steam. The degasser also serves to concentrate the C₄acetylenes (vinyl- and ethyl-acetylene), 1,2-butadiene and VCH. Thesecomponents are removed at their point of highest concentration via aliquid side draw.

A comparison of flow rates and energy requirements for these processesis presented in Table 1.

TABLE 1 Example 1 Conventional Compressorless Duty Duty Duty UnitDegassing mm Temp. Degassing mm Temp. Degassing mm Temp. Operation kg/hkcal/h ° C. kg/h kcal/h ° C. kg/h kcal/h ° C. Rectifier 21797 4.361  9039918 9.588 120/103.8 27834  7.917 120/108 Reboiler Degasser 26429 9.678137.9/129 — — — — — — Feed Flash Drum 26429 0.725 — — — — — — Degasser6326 —  129/124.6 22101 1.163 120/104.5 16980 9.861 + 0.730 160/147Inlet Flash Degasser 9945 10.757 175 6793 9.168 150 9835 11.053 193Total 64498 25.521 68812 19.919  54.649 29.561 Solvent Flow 290082311433 316012 (kg/h) Compressor 18703 31256 — Flow (kg/h) Compression1.61 3.42 — Ratio Power (kw) 292.6 799.2 — Rich Solvent 64475 6878354650 Dissolved Gas (kg/h)

As shown by the table above, for the process of Example 1, process heat(degasser bottoms) is advantageously exchanged for high pressuredegassing (rectifier bottoms) and advantageously exchanged for lowpressure degassing (degasser feed). Process heat is added to therectifier bottoms (high pressure level) in two steps vs. the one-stepheat addition in the conventional design. The first step is a “mild”heat input via the Rectifier reboiler, which is a once-thru reboilerthat heats the rich solvent from 80° C. to 90° C. (only 10° C. delta T).At this low outlet temperature, a suppressed-vaporization reboiler isnot required, thus no rectifier bottoms pump is required. Even thoughconditions are mild, one-third of the dissolved gases are removed fromthe rich solvent in this first degassing step. The percent vaporizationin the rectifier reboiler is also quite low (6.2 wt. %). The heatingmedium inlet temperature on the hot side of the rectifier reboiler iscontrolled at 120° C., which is significantly below the 150° C. heatingmedium used in the conventional design. The combination of lowtemperatures (both hot side and cold side) and the low percentvaporization avoids the issue of fouling in the rectifier reboiler.

In contrast, the conventional design heats up the rectifier bottoms from76° C. to 120° C. (44° C. delta T) in the first degassing step. Theprocess of Example 1 achieves almost half of the degassing as in theconventional design, with only a 10° C. temperature rise vs. the 44° C.temperature rise of the conventional design. Thus, there are clearlydiminishing returns when trying to accomplish the high pressuredegassing in a single step, as shown in FIG. 5.

The second step is a more “severe” heat input via the degasserfeed/effluent exchanger, which requires a suppressed-vaporizationreboiler (similar to the conventional design). In the process of Example1, the rich solvent is heated to about 140° C. in a suppressedvaporization heater, the degasser feed/effluent exchanger, and flashedinto the feed drum located at the top of (and part of) the degasser. Thedrum operates at a pressure slightly above the rectifier bottomspressure, so the vapor flows freely back to the rectifier bottomswithout the need for vapor recompression. The high level heat input isin the form of the flashed vapor returned to the rectifier. Theun-flashed liquid then flows by pressure differential and gravity to thetop of the degasser, where additional flashing occurs.

The two-step high pressure degassing is also more efficient in terms ofC₄ degassing. Two stages of flashing provide better separation of C₄sfrom the solvent. In other words, more C₄s and less solvent arevaporized compared to the conventional design.

Because of the higher operating pressure of the degasser, recycle gas isreturned to the Rectifier at a slightly higher temperature than in theconventional design and more “pre-degassing” occurs in the bottomsection of the rectifier. This contributes to an overall lower degassingrequirement.

Process heat is added to the degasser feed (low pressure degassing) bymeans of the sensible heat contained in the un-flashed liquid in thedegasser feed drum. In other words, part of the heat input from thedegasser feed/effluent exchanger ends up in the feed to the degasser.

Because almost 75% of the dissolved gases in the rich solvent aredegassed at the higher pressure level (not requiring recompression) vs.only 58% in the conventional design, the capacity of the liquid ringcompressor of Example 1 is only 60% of the capacity of the screw-typecompressor required for the conventional design. The degasser operatesat a pressure slightly less than 2 kg/cm² below the rectifier pressure.Consequently, the compression ratio required between the degasser andthe rectifier is only 1.61 vs. 3.42 for the conventional design. Thecombination of lower flow and lower compression ratio allows the use ofa liquid ring compressor instead of the more-expensive centrifugal orscrew-type compressor employed in the conventional design. The smallersize (flow and compression ratio) of the liquid ring compressor makes iteven less expensive.

The combination of lower flow and lower compression ratio results in apower consumption that is only 37% of the power required in theconventional design, even after accounting for the expected loweradiabatic efficiency of the liquid ring compressor (50% vs. 76%).

The cooling column used in the conventional design is eliminated in theprocess of Example 1, and its function is largely replaced by the liquidring compressor. Thus, the liquid ring compressor accomplishes twooperations: compressing the degasser overhead; and cooling thecompressed gas.

As noted above, in some embodiments, a second degasser feed heater canbe added at the bottoms of the degasser feed drum to provide someadditional low-level utility heat to the degasser feed. In this case,the heater could be located at grade and the liquid static head at theinlet to the exchanger would be used to suppress vaporization. The lowpressure steam utilized would displace an equivalent amount of mediumpressure steam, resulting in better economics. This option may alsodepend on project specific availability of low pressure steam andrelative utility costs.

The combined degasser feed drum/degasser design has no significant costpenalty. For example, the degasser feed/effluent exchanger is asuppressed vaporization reboiler, there is already sufficient fluidpressure to overcome the static head to feed into the drum located atthe higher elevation. In the conventional design this is simply chewedup across the control valve. Thus, there is no extra cost associatedwith pumping into the drum mounted on top of the degasser. Additionally,the degasser is designed to be liquid-filled during chemical cleaningand passivation. The addition of a drum on the top of the degasser doesnot add significantly to the cost of the tower. Furthermore, until justrecently, all previous degasser designs had the acetylene washer mountedon the side of the degasser. The acetylene washer is significantlybigger and heavier than the feed drum, and it was installed in anasymmetric position. There is little or no extra cost associated withmounting the drum on top of the degasser. Further, the cost of the drumon top of the tower is less than for a stand-alone drum: only oneadditional head is required; the incremental cost for additional shelllength is small; there is no additional piece count; and no additionalplot area is required.

The degasser operates at a higher pressure and temperature than thedegasser in the conventional design. Although the bottoms temperature(175° C.) is higher than the conventional design (150° C.), it issignificantly lower than the degasser bottoms temperature in thecompressorless design (193° C.). Thus, the process of Example 1 benefitsfrom the pre-degassing in the bottom of the rectifier and the degassingand compression area.

Compared to the conventional butadiene extraction design, embodimentsdisclosed herein may have one or more of the following advantages. 1.Operation of the degasser at a higher pressure and temperature. 2.Replacement of the conventional recycle gas compressor (centrifugal orscrew-type) with a smaller, less-expensive liquid ring compressor. 3.Replacement of the conventional solvent exchangers (3-shell design) withthe following: a. rectifier reboiler (1 shell) b. degasser feed/effluentexchanger (2 shells). 4. Use of a feed flash drum mounted on top of thedegasser. This allows the recovery of flashed vapor without the need forrecompression and also eliminates the need for a second set of pumpsbetween the rectifier and the degasser (see item 4). 5. Elimination ofthe rectifier bottoms pumps (high capacity/high head). 6. Elimination ofcooling column (cooling provided in the liquid ring compressor). 7.Elimination of the cooling column bottoms pumps (coolant flow is smalland by pressure letdown). 8. Smaller solvent and water cooler (only 1shell). 9. Lower equipment cost. 10. Lower operating cost.

Compared to the Compressorless butadiene extraction design, embodimentsdisclosed herein may have one or more of the following advantages. 1.Operation of the degasser at a lower pressure and temperature. 2.Replacement of the solvent exchangers (3-shell design) with a smallerrectifier reboiler (1-shell design). 3. Smaller degasser feed/effluentexchanger. 4. Addition of a small, low-cost liquid ring compressor,cooler and knock-out drum. 5. Addition of a feed flash drum mounted ontop of the degasser. 6. elimination of the rectifier bottoms pumps (highcapacity/high head). 7. Significantly less risk: degasser bottomstemperature is 175° C. vs. 193° C. for compressorless option, resultingin less fouling and a higher limit on inlet C₄ acetylenes concentration.8. Higher yield: degasser bottoms temperature is 175° C. vs. 193° C. forcompressorless option, resulting in less fouling and a higher limit oninlet C₄ acetylenes concentration. 9. Expected lower equipment cost. 10.Expected lower operating cost.

As shown above, butadiene extraction processes according to embodimentsdisclosed herein may be operated at a relative intermediate pressureusing a liquid ring type compressor. The use of a liquid ringcompressor, among other process options presented herein, mayadvantageously reduce capital and operating costs, similar to thecompressorless option, while mitigating the risks associated with thehigher operating temperatures and pressures of the compressorlessoption. Thus, the embodiments disclosed herein encompass the bestfeatures of the conventional design (low pressure, with a compressor)with the advantages of the compressorless design (low capital andoperating cost), as well as other advantages unique to the systemsdisclosed herein.

While the disclosure includes a limited number of embodiments, thoseskilled in the art, having benefit of this disclosure, will appreciatethat other embodiments may be devised which do not depart from the scopeof the present disclosure. Accordingly, the scope should be limited onlyby the attached claims.

What is claimed:
 1. A process for recovering 1,3-butadiene from a C₄fraction, comprising: feeding a hydrocarbon fraction containing butanes,butenes, 1,2-butadiene, 1,3-butadiene, C₄ acetylenes, C₃ acetylenes, andC₅₊ hydrocarbons to an extractive distillation system; contacting thehydrocarbon fraction with a solvent in the extractive distillationsystem to selectively dissolve a portion of the hydrocarbon fraction;recovering a vapor fraction comprising a first portion of the butanesand the butenes from the extractive distillation system; recovering anenriched solvent fraction comprising the 1,3-butadiene, the1,2-butadiene, C₄ acetylenes, C₃ acetylenes, C₅₊ hydrocarbons, and asecond portion of the butanes and the butenes; feeding the enrichedsolvent fraction to a rectifier to at least partially degas the enrichedsolvent; recovering a second portion of the butanes and butenes from therectifier as an overheads fraction; recovering the C₃ and C₄ acetylenes,1,3-butadiene, 1,2-butadiene, and C₅₊ hydrocarbons from the rectifier asa side draw fraction; recovering a partially degassed solvent comprising1,2-butadiene and C₄ acetylenes from the rectifier as a bottomsfraction; feeding at least a portion of the partially degassed solventto a degasser to further degas the solvent; recovering an overheadsfraction comprising at least one of C₄ acetylenes and 1,2-butadiene fromthe degasser; recovering a side draw fraction comprising the C₄acetylenes from the degasser; recovering a bottoms fraction comprisingdegassed solvent from the degasser; compressing the degasser overheadsfraction using a liquid ring compressor; and recycling at least aportion of the compressed degasser overheads fraction to the rectifier.2. The process of claim 1, further comprising phase separating thecompressed degasser overheads fraction to recover a condensate fractionand recycling at least a portion of the condensate fraction to theliquid ring compressor.
 3. The process of claim 1, further comprising:heating the partially degassed solvent via indirect heat exchange tovaporize at least a portion of the dissolved 1,2-butadiene and/or C₄acetylenes; feeding the heated partially degassed solvent to a degasserfeed drum to phase separate the vaporized portion of the 1,2-butadieneand/or C₄ acetylenes from the heated partially degassed solvent;recovering a vapor phase from the degasser feed drum comprising at leastone of 1,2-butadiene and C₄ acetylenes; recovering a liquid phase fromthe degasser feed drum; feeding the liquid phase from the degasser feeddrum to the degasser as the at least a portion of the partially degassedsolvent.
 4. The process of claim 3, wherein the degasser feed drum andthe degasser are integral.
 5. The process of claim 3, further comprisingadmixing the vapor phase recovered from the degasser feed drum and theat least a portion of the compressed degasser overheads and recyclingthe combined portion to the rectifier.
 6. The process of claim 5,wherein a ratio of the at least a portion of the compressed degasseroverheads to the vapor phase is in a range from about 0.1:1 to 1:1. 7.The process of claim 3, wherein the degasser feed drum operates at apressure in the range from about 3.5 to about 5.5 kg/cm² gage and atemperature in the range from about 120° C. to about 140° C.
 8. Theprocess of claim 7, wherein the degasser pressure is greater than anoperating pressure of the rectifier.
 9. The process of claim 3, whereinthe heating the partially degassed solvent via indirect heat exchangecomprises at least one of: contacting the partially degassed solventwith the degasser bottoms fraction via indirect heat exchange; andcontacting the partially degassed solvent with a heat exchange mediumcomprising at least one of a synthetic organic heat transfer fluid,water, and steam.
 10. The process of claim 9, wherein a heatexchanger(s) for contacting the partially degassed solvent comprises asuppressed vaporization heater.
 11. The process of claim 1, wherein thedegasser operates at an overhead pressure in the range from about 1.5 toabout 3.5 kg/cm² gage and an overhead temperature in the range fromabout 110° C. to about 140° C.
 12. The process of claim 11, wherein thedegasser operates at a bottoms pressure in the range from about 160° C.to about 190° C.
 13. The process of claim 1, wherein the rectifieroperates at an overhead pressure in the range from about 3 to about 5kg/cm² gage and an overhead temperature in the range from about 50° C.to about 70° C.
 14. The process of claim 13, wherein the rectifieroperates at a bottoms temperature in the range from about 70° C. toabout 100° C.
 15. The process of claim 14, wherein heat is supplied tothe rectifier via indirect heat exchange in a reboiler using a heatingmedium having a temperature of less than 130° C.
 16. The process ofclaim 15, wherein the reboiler comprises a single pass heat exchanger.17. The process of claim 16, wherein a process-side temperature increaseacross the reboiler is in the range from about 5° C. to about 15° C. 18.The process of claim 17, wherein a percent vaporization across thereboiler is in the range from about 3 wt. % to about 9 wt. %.